JP5185623B2 - Continuous process for the production of phenol from benzene in a fixed bed reactor. - Google Patents

Description

Detailed Description of the Invention

(Technical field)
The present invention relates to a continuous process for the production of phenol by direct oxidation of benzene with hydrogen peroxide carried out in a fixed bed reactor.
More particularly, the present invention relates to a process for producing phenol in which the oxidation reaction is carried out in the presence of a catalyst based on titanium silicalite and under specific operating conditions.
The present invention also relates to a method for activating a catalyst based on titanium silicalite TS-1.

(Background technology)
Phenol is a very important industrial intermediate and is widely used in the production of polycarbonates or other phenolic resins.
Phenol is currently produced by the “Hock process”, which is intended for alkylation of benzene to cumene and subsequent oxidation of cumene to hydroperoxide (which decomposes to phenol and acetone). doing.
Various processes for the production of phenol based on the direct oxidation of benzene with hydrogen peroxide in the presence of a suitable catalyst system are known in the art.
For example, methods are known which are carried out in the presence of a catalyst based on titanium silicalite and in an organic solvent capable of promoting contact between the organic substrate and hydrogen peroxide.
The conversion rate and selectivity of the production process of phenol by direct oxidation can be improved, for example, by operating in the presence of certain solvents such as sulfolane (EP A 919531).
In this case, the above process is carried out in a batch type reactor operating in a two-phase reaction system consisting of a solid catalyst and an organic phase comprising sulfolane / water / benzene in such a ratio that makes the reaction mixture homogeneous. .
An improvement in the productivity of the phenol production process can also be obtained by activation of the catalyst with hydrogen peroxide and fluoride ions as described in European patent application EP A 958861.
Further improvements can be obtained by operating in a three-phase reaction system consisting of a solid catalyst, an aqueous phase and an organic phase comprising an aromatic compound and a solvent, as described in patent application PCT / EP02 / 12169. it can.
The process for the production of phenol by direct oxidation of benzene described in the above-mentioned prior art is generally performed in a CSTR type reactor in which a catalyst based on titanium silicalite is maintained in suspension in the form of a fine powder. We are carrying out.
In these processes, the separation of the reaction effluent from the catalyst can only be effected effectively by filtration in the reactor.
However, this operation cannot be performed in the prior art methods because the filtration surface is necessarily too high as a result of the low concentration of product.
As a result, the only possible technical solution consists of performing the filtration outside the CSTR reactor. However, this solution is technically complex because it requires the use of both additional equipment and wear resistant materials and also the maintenance of the effectiveness of the filter.

(Disclosure of the Invention)
We have now found a method that enables the direct synthesis of phenol from benzene with high conversion and productivity, while at the same time eliminating the above problems by filtration.
In fact, the present invention operates the above process in a fixed bed reactor under specific operating conditions and according to two different procedures, i.e. both in a homogeneous liquid phase or in two liquid phases. It is intended to be carried out continuously using sulfolane as.
In accordance with the above, the object of the present invention is a continuous process for the production of phenol by direct oxidation of benzene with hydrogen peroxide in the presence of a catalyst based on titanium silicalite TS-1, characterized in that it comprises the following steps: About:
(a) performing the above process in a fixed bed reactor containing a TS-1 based catalyst at a temperature of 80-120 ° C. and a pressure of 3-15 atm;
(b) The reactor is fed with a stream containing H ₂ O ₂ , benzene, sulfolane and water in a single phase or a double phase, and the amount of each single component is 100 units each. Each feed is in the range of 0.2-6, 15-60, 30-80, 0.5-30 parts by mass, and their total flow rate is such that the residence time in the reactor is in the range of 0.3-2 minutes Calculating step (residence time is the ratio of catalyst mass to feed flow rate);
(c) recovering product, by-products and solvent from the liquid stream exiting the reactor.
The method of the invention makes it possible to obtain various advantages:
The process according to the invention no longer requires filtration, and also avoids recirculation of part of the liquid reaction effluent, thus greatly simplifying the operations involved in the oxidation process of the aromatic substrate;
The method of the invention saves on investment costs for plant production;
-The method of the present invention is more versatile because it also allows operation in two liquid phases (three phase systems).

The process of the present invention can be carried out efficiently in succession and using several reactors.
In this case, it has been found that the selectivity can be improved by dividing the supply of hydrogen peroxide equally between the respective reactors.
The process of the invention is carried out in a fixed bed reactor (plug-flow type) loaded with a catalyst based on titanium silicalite.
The catalyst of the present invention is titanium silicalite having an MFI structure and having the general formula xTiO _2. (1-x) SiO ₂ , where x is in the range of 0.0001 to 0.04.
The structural features of these catalysts and their method of preparation are described in US Pat. No. 4,410,501, US Pat. No. 4,954,653, US Pat. No. 4,701,428, EP A 906784.
The catalyst of the present invention is used in an extruded form by blending titanium silicalite with an inert ligand such as silica according to methods known to those skilled in the art.
Examples of these formulations can be found in US Pat. No. 6,491,861, US Pat. No. 6,551,546, US 2003/0078160, US 2003/0130116.

The reaction in the production of phenol by direct oxidation of benzene with H ₂ O ₂ follows two different procedures: in a homogeneous liquid phase (dual phase system) or using two liquid phases (triple phase), both using sulfolane as solvent. System).
In the phenol production process, the feed ratio of H ₂ O ₂ , benzene, sulfolane and water results in the presence of a single liquid phase or two immiscible phases.
When operating in a single liquid phase, the feed rates of H ₂ O ₂ , benzene, sulfolane and water are generally 0.2-4, 20-60, 40-80, 0.5- While operating in the presence of two immiscible phases, these amounts range from 0.3 to 6, 15 to 50, 30 to 70, 10 to 30 parts by weight. Is in.
In both cases, the total flow rate is calculated such that the residence time is in the range of 0.3-2 minutes.
The oxidation of benzene to phenol with hydrogen peroxide is preferably carried out in an apparatus consisting of a series of fixed bed plug flow reactors (minimum 2 to maximum 10 reactors).
Each reactor consists of a single catalyst bed with or without an external heating jacket.
In this latter case, each reactor operates adiabatically and the catalyst activation process takes place in a specific apparatus.
The basic structure of the reaction zone shown in FIG. 1 takes into account, inter alia, six tubular reactors designated as R110 ,..., R160 . The liquid flow shown is intended for each case to be a homogeneous liquid phase or a total of two immiscible phases.
All the benzene and sulfolane required for the reaction are mixed (M110) and sent to the first reactor (R110).
On the other hand, the hydrogen peroxide in the aqueous solution is divided into equal parts and added to the mainstream in front of each reactor (in M110 to M160) to keep the concentration low and consequently suppress its decomposition into oxygen To do.
Each reactor completely consumes the hydrogen peroxide feed.
By subdividing the hydrogen peroxide between each reactor, the concentration of the phenol formed in the first reactor is also kept low; in this method, fewer reactions result in the formation of by-products (catechol and hydroquinone). It is possible to suppress more efficiently than when operating with several reactors.

The amount of reagent fed to each reactor will vary depending on whether the procedure is performed in single phase or dual phase.
Typically, when operating in a single layer, the amount of each component of the feed stream (expressed in parts by weight per 100 unit feed) is 0.2-4 for hydrogen peroxide, 20-60 for benzene, 40-40 for sulfolane. 80, 0.5-5 in water; preferably 0.2-4 in hydrogen peroxide, 30-50 in benzene, 45-70 in sulfolane, 0.5-5 in water.
On the other hand, to operate in a dual phase, the amount of each component of the feed stream (expressed in parts by mass per 100 unit feed) is 0.3-6 for hydrogen peroxide, 15-50 for benzene, 30-70 for sulfolane. 10 to 30 in water; preferably 0.3 to 5 in hydrogen peroxide, 20 to 40 in benzene, 35 to 60 in sulfolane, and 10 to 30 in water.
Nitrogen is sent to the top of each reactor to keep the vapor outside the explosion limit.
The reactor can be operated either under adiabatic conditions or in the presence of thermostat regulation, with a temperature range of 80-120 ° C. and a pressure of 3-15 atm, preferably 90-110 ° C. and 5-7 atm.
Heat exchangers ( E 110,..., E 160) exist between the reactors to lower the temperature of the effluent before feeding to the next reactor.
In the case of an adiabatic reactor, the absence of heat exchange (in the reactor) makes the reactor particularly simple in terms of construction.
The total feed flow rate is calculated so that the residence time in the reactor is in the range of 0.3 to 2 minutes.
By operating according to the invention described so far, 100% conversion of H ₂ O ₂ has been obtained with a selectivity of benzene of around 85-92%.
Furthermore, by operating in the dual phase, the reaction rate is increased compared to the homogeneous phase, which allows the use of lower amounts of catalyst and smaller reactors, and also The temperature is controlled to a lower value.
The dual phase makes it possible to keep the amount of catechol and hydroquinone forming benzene selectivity also constant.

The catalytic properties of the catalyst used in the oxidation of benzene can be improved by activation treatment with fluoride and hydrogen peroxide.
In a batch process, the treatment can be carried out on the powder before extrusion, and after treatment, a series of water washings and filtration (as described in patent EP 958861).
It has now been found that the activation step can be carried out directly and continuously in the column where the synthesis reaction takes place: In this method, the treatment is simpler and easier to apply on an industrial scale.
The activation method of the present invention comprises treating the amount of catalyst and reagent to be treated with H ₂ O _{2 in} the range of 0.5-3.0, preferably 2.5 fluorine / titanium molar ratio and 3.0-15, preferably 11 It is intended to be selected to have a titanium / titanium molar ratio.
The catalyst can be used as such or intimately mixed with an inert material of similar dimensions in an amount equal to the mass of the catalyst itself.
The inert material may be selected from quartz, corundum, ceramic material, glass, extruded silica, preferably quartz.
A catalyst or mixture of its inerts is charged into the reactor and then brought to a temperature in the range of 20-120 ° C, preferably 80 ° C, and ammonium acid fluoride (NH ₄ HF ₂ ) and Supply an aqueous solution of hydrogen peroxide.
The supply time is in the range of 2-6 hours, preferably 4 hours.
The feed solution comprises ammonium acid fluoride in a concentration range of 0.1% to 1% by weight, preferably 0.25% by weight, and peroxidation in a range of 3% to 10% by weight, preferably 4.8% by weight. Contains hydrogen. At the end of the reaction, water is supplied to the reactor to remove the reagent residue, and then the liquid is discharged and the catalyst is subjected to calcination or drying treatment.

The liquid effluent leaving the reactor at the end of the reaction containing products and by-products and the solvent to be recovered for reuse must be subjected to a purification zone.
The purification of the product and the recovery of the solvent can be carried out by the method described in patent EP 3076502.
However, this method has the disadvantage that the solution coming out of the reactor and sent to the distillation zone contains high amounts of water and sulfolane, and their presence makes it difficult to produce pure phenol.
We have now found a method for recovering the product and solvent that allows for a reduction in the amount of water and sulfolane sent to the distillation, thus resulting in a more efficient and economical distillation operation.
Indeed, the process of the present invention contemplates subjecting the two-phase liquid stream exiting the reactor to liquid-liquid extraction using water and benzene as extraction solvents.
Water is sent to the column head to function to properly extract sulfolane, while benzene is sent to the bottom of the column to function to properly extract phenol.
Biphenols and other by-products are divided into two streams.
An organic stream rich in benzene and phenol is sent to distillation, while an aqueous stream rich in water and sulfolane is sent directly to chlorination to recover biphenols.
More specifically, a reaction mixture containing benzene, water, phenol, sulfolane and reaction by-products (biphenols) is sent to one or more extraction towers, which also send benzene and water. Water (fresh or recycled from other equipment) is sent from the top, the reaction mixture is sent to the midpoint, while benzene (fresh or recycled from other equipment) is sent from the bottom. The light organic phase, rich in benzene and phenol as opposed to the mixture coming out of the reaction and free of much sulfolane and water, exits from the head of the tower; whereas it is rich in water and sulfolane as compared to the mixture coming out of the reaction. A heavy aqueous phase poor in benzene and phenol exits from the bottom of the tower.
The organic phase is sent to a distillation zone while the aqueous phase is sent to a biphenol chlorination unit.
The distillation and biphenol recovery zone has a shape similar to that described in US Patent No. 10/716460, but has reduced size and lower energy consumption due to the previous extraction zone.
Besides the separation and purification of phenol, the above procedure used makes it possible to obtain a purification solvent containing benzene necessary for recycling to the oxidation reactor, and furthermore, biphenols dissolved in water are It is then reconverted to phenol by catalytic hydrodeoxygenation as described in US Patent No. 10/716460.

Example 1
Catalyst activation
7100 kg of TS-1 catalyst (extruded with 15% by weight silica) is charged into a tubular reactor as described in Examples 2 and 3 at a temperature of 80 ° C.
207 kg of acidic ammonium fluoride (NH ₄ HF ₂ ) and 3944 kg of 30% by mass hydrogen peroxide are dissolved in 81,000 kg of distilled water and fed to the reactor at a flow rate of 20 m ³ / h. At the end of the addition, 40,500 kg of water is fed to the reactor to remove reagent residues. Finally, the reactor is emptied and the temperature is brought to 350 ° C. for 8 hours under a stream of inert gas.
The catalyst thus treated is ready for phenol synthesis and has been found to have the following composition: SiO ₂ = 98.20%, TiO ₂ = 1.80%.

Example 2
Production of phenol in homogeneous phase
The benzene and sulfolane supplied to M110 are 251.1 t / h and 363.35 t / h, respectively. Each of the six streams of hydrogen peroxide (40% by weight of water) is equal to 3.36 t / h.
M110 mixes the feed and E110 heats the feed from room temperature to 100 ° C.
In R110, the hydrogen peroxide is completely converted into the reaction product, and the temperature rises from 100 ° C. to 112 ° C. by adiabatic operation.
M120 mixes the first reactor effluent with a second amount of hydrogen peroxide and E120 cools this stream to 100 ° C. Since the amount of hydrogen peroxide is about the same, the temperature rise at R120 also goes from 100 ° C to 112 ° C.
Subsequent M130, ..., M160 and E130, ..., E160 have the same function as M120 and E120; reactors R130, ..., R160 behave similarly to R120.
Each of the reactors contains 9,000 kg of catalyst (activated as described in Example 1) and has the following dimensions: D = 2.75 m, H = 3.14 m.
The liquid effluent leaving the last reactor has a flow rate of 639.8 t / h and the following composition: benzene = 35.9% by weight, phenol = 3.20% by weight, catechol = 0.55% by weight, hydroquinone = 0.29% by weight, sulfolane = 56.79 wt%, water = 3.23 wt%.
The overall performance obtained is as follows:
H ₂ O ₂ conversion rate (%): 100
Benzene selectivity (%): 87.5
H ₂ O ₂ selectivity (%): 74.0
H ₂ O ₂ to O ₂ (%): 5.3
From H ₂ O ₂ to (catechol + hydroquinone) (%): 21.0

Example 3
Production of phenol in the double phase
The benzene, water and sulfolane fed to M110 are equal to 143.19 t / h, 2,226.33 t / h and 96.36 t / h, respectively. Each of the six streams of hydrogen peroxide (65% by weight of water) is equal to 6.12 t / h.
M110 mixes the feed and E110 heats the feed from room temperature to 95 ° C.
In R110, the hydrogen peroxide is completely converted into the reaction product, and the temperature rises from 95 ° C. to 108.5 ° C. by adiabatic operation.
M120 mixes the first reactor effluent (two-phase mixture) with a second amount of hydrogen peroxide and E120 cools this stream to 95 ° C. Since the amount of hydrogen peroxide is comparable, the temperature rise at R120 is also 95 ° C to 108.5 ° C.
Subsequent M130, ..., M160 and E130, ..., E160 have the same function as M120 and E120; reactors R130, ..., R160 behave similarly to R120.
Each of the reactors contains 7,100 kg of catalyst (activated as described in Example 1) and has the following dimensions: D = 2.54 m, H = 2.91 m.
The liquid stream leaving the last reactor (two-phase) has a flow rate of 493.80 t / h and the following composition: benzene = 23.73 wt%, phenol = 4.41 wt%, catechol = 0.54 wt%, hydroquinone = 0.29 wt% , Sulfolane = 45.83 wt%, water = 24.73 wt%.
The overall performance obtained is as follows:
H ₂ O ₂ conversion rate (%): 100
Benzene selectivity (%): 85
H ₂ O ₂ selectivity (%): 61.5
H ₂ O ₂ to O ₂ (%): 16
From H ₂ O ₂ to (catechol + hydroquinone) (%): 19.8

1 is a schematic diagram showing the basic structure of a reactor for carrying out the method of the present invention.

Explanation of symbols

R110 ~ R160 reactor
E110 to E160 heat exchanger
M110-M160 mixer

Claims (12)

A process for the continuous production of phenol by the direct oxidation of benzene with hydrogen peroxide in the presence of titanium silicalite TS-1 catalyst characterized in that it comprises the following steps (a) to (c):
(a) the direct oxidation is carried out at a temperature of 80 to 120 ° C. and a pressure of 3 to 15 atm in the same fixed bed reactor containing the catalyst and used for the activation of the catalyst. Process;
(b) A stream containing H ₂ O ₂ , benzene, sulfolane and water in the range of 0.2-6, 15-60, 30-80, 0.5-30% by mass, respectively, is supplied to the reactor , and Obtaining a single liquid phase or two immiscible liquid phases, wherein the residence time of the stream in the reactor is in the range of 0.3 to 2 minutes;
(c) recovering product, by-products and solvent from the stream exiting the reactor.   The process according to claim 1, wherein the direct oxidation is carried out at a temperature in the range of 90-110 ° C and a pressure of 5-7 atm. The stream containing H ₂ O ₂ , benzene, sulfolane and water is in two immiscible liquid phases and the residence time of the stream in the reactor ranges from 0.3 to 2 minutes. The method described.   The process according to claim 1, wherein the direct oxidation is carried out in an apparatus consisting of 2 to 10 fixed bed plug flow reactors arranged in series.   The method of claim 4, wherein the hydrogen peroxide feed is divided equally between each reactor and added to the stream.   The process of claim 4 wherein the residence time of the stream in each reactor is in the range of 0.4 to 2.0 minutes.   The process of claim 1 wherein the catalyst is used in an extruded form. The titanium silicalite TS-1 catalyst used in the direct oxidation is activated in a step including the following steps (d) to (f) in the same fixed bed reactor as the fixed bed reactor in the step (a). The method of claim 1 wherein:
(d) In the reactor containing the catalyst as it is or mixed with an inert material selected from quartz, corundum, glass, extruded silica, at a temperature of 20-120 ° C., in the range of 0.1% to 1% by weight Supplying an aqueous solution containing NH ₄ HF _{2 at} a concentration of ₅ % and hydrogen peroxide at a concentration ranging from 3% to 10% by weight for 2 to 6 hours;
(e) supplying water to the reactor at the end of the reaction;
(f) A step of drying or calcining the catalyst contained in the reactor. The method according to claim 8, wherein in the step (d), an aqueous solution containing NH ₄ HF _{2 having} a concentration of 0.25% and hydrogen peroxide having a concentration of 4.8% is supplied to the reactor at a temperature of 80 ° C. . In the step (d), the aqueous solution containing NH ₄ HF ₂ and hydrogen peroxide has a molar ratio of fluorine contained in the aqueous solution and titanium contained in the catalyst in the range of 0.5 to 3.0, and supplying an amount such that the molar ratio of H ₂ O ₂ and titanium is in the range of 3.0 to 15, the method of claim 8.   The method according to claim 10, wherein a molar ratio of the fluorine and titanium contained in the aqueous solution is 2.5.   The method of claim 8, wherein the inert material is quartz.